Production of terephthalic acid

ABSTRACT

An aromatic carboxylic acid such as terephthalic acid is produced by the liquid phase oxidation of a precursor thereof, the oxidation being carried out in such a way that substantially all of the aromatic carboxylic acid produced in the course of the reaction is maintained in solution during the reaction.

This application claims benefit of provisional application Ser. No.60/039,662 filed Feb. 28, 1997.

This invention relates to the production of aromatic carboxylic acidswhich are sparingly soluble in acetic acid and water, particularlyterephthalic acid.

Terephthalic acid is an important intermediate for the production ofpolyester polymers which are used typically for fibre production and inthe manufacture of bottles. Current state-of-the-art technology for themanufacture of terephthalic acid involves the liquid phase oxidation ofparaxylene feedstock using molecular oxygen in a solvent comprisinglower C2 to C6 aliphatic monocarboxylic acid, usually acetic acid, inthe presence of a dissolved heavy metal catalyst system incorporating apromoter such as bromine. The reaction is carried out in at least onestirred vessel under elevated temperature and pressure conditions,typically 150 to 250° C. and 6 to 30 bara respectively, with air beingsparged into the reaction mixture and typically produces terephthalicacid in high yield, e.g. at least 95%. Isothermal reaction conditionsare maintained in the oxidation vessel by allowing evaporation of thesolvent, together with water produced in the reaction, the resultingvapour being condensed and returned to the reactor vessel as reflux. Inthe conventional production of terephthalic acid, because terephthalicacid is only sparingly soluble in the solvent, a substantial proportionthe product precipitates in the course of the reaction and as a resultimpurities such as 4-carboxybenzaldehyde (4-CBA) and colour bodiesco-precipitate with the terephthalic acid to produce a crude productwhich, to meet the requirements of many polyester producers, has to bepurified to reduce its impurity content. In one purification process,the crude product is dissolved in water and, under elevated temperatureand pressure conditions, is contacted with hydrogen in the presence of ahydrogenation catalyst, the purified terephthalic acid thereafter beingrecovered by crystallisation and solids-liquid separation techniques.

The present invention seeks to provide a process for the production ofterephthalic acid in such a way as to afford scope for achieving asufficiently pure product for subsequent use without necessarily havingto carry out an additional purification process.

According to a first aspect of the present invention there is provided aprocess for the production of terephthalic acid by the liquid phaseoxidation of a precursor of terephthalic acid with oxygen in a reactionmedium containing the precursor and a solvent under conditions such thatsubstantially all of the terephthalic acid produced in the oxidationreaction zone is maintained in solution during the reaction,characterised in that the oxidation reaction is carried out by passingthe reaction medium through the reaction zone in a continuous plug flowreaction regime.

Preferably the reactor is a plug flow reactor or a series of two or moreplug flow reactors, preferably operated in a non-boiling mode, althoughthe various aspects ot the invention defined herein are not limited tothis particular type of continuous flow reactor. For instance, thereaction may be carried out in a series of non-boiling continuousstirred tank reactors so as to approximate a continuous plug flow regimeor in a reaction system comprising one-or more non-boiling continuousstirred tank reactors and one or more plug flow reactors arranged in anysequence.

By “continuous plug flow regime” we mean a reactor in which reactantsare introduced and products withdrawn simultaneously in a continuousmanner, as opposed to a batch-type reactor. The residence time of thereaction medium within the reaction zone is generally no more than 10minutes and is preferably no more than 8 minutes, with residence timesof 5 minutes or less, e.g. 3 minutes or less, being achievable.

According to a second aspect of the present invention, which may be butis not necessarily used in conjunction with the first aspect of theinvention, there is provided a process for the production ofterephthalic acid by the liquid phase oxidation of a precursor ofterephthalic acid with oxygen in a reaction medium containing theprecursor and a solvent under conditions such that substantially all ofthe terephthalic acid produced in the oxidation reaction zone ismaintained in solution during the reaction, characterised in that theoxidation reaction is carried out with substantially all of the oxygendissolved in the reaction medium.

Thus, in this aspect of the present invention, the liquid phaseoxidation reaction is carried out in such a way as to maintainsubstantially all of the resulting terephthalic acid in solution duringthe reaction thereby reducing the extent to which the main impurity,4-CBA, contaminates the recovered terephthalic acid as a result ofco-precipitation during the reaction. Substantially all of the oxygenutilised in the process according to this aspect of the presentinvention is dissolved in the reaction medium. The use of dissolvedoxygen in the reaction medium allows the oxygen to be more uniformlydistributed throughout the reaction medium. In this manner, oxygenstarved regions within the reaction medium can be minimised withconsequential reduction in the formation of undesirable reactionby-products such as trimellitic acid, benzoic acid and colour bodies.Overall, this leads to the possibility of producing product with a lowlevel of contamination and without undue solvent burning which, in turn,allows elimination of the purification process conventionally employedin the production of terephthalic acid of sufficient quality for use inhigh grade polyester manufacture.

Although in the above aspect and other aspects of the inventiondisclosed herein, it is preferred that all of the terephthalic acidproduced in the reaction is maintained in solution during the reaction,we do not exclude the possibility of some precipitation during thereaction, e.g. up to 10%, more usually no more than 5% but desirably nomore than about 2% by weight of the terephthalic acid produced mayprecipitate during the course of the reaction.

Preferably the reaction medium is produced by combining at least twoseparate liquid phase components and at least part of the oxygen isadded to and dissolved in one or more of said liquid phase componentsbefore such components are combined to form the reaction medium.

For instance, the separate liquid phase components may include onecomponent consisting of or containing said precursor and a secondcomponent consisting of or containing said solvent and at least part ofthe oxygen required for the reaction may be added to and dissolved inthe second component so that reaction between the oxygen and theprecursor cannot commence until the components are combined to form thereaction medium.

Usually the solvent is predominantly an aliphatic monocarboxylic acid(preferably containing 2 to 6 carbon atoms) and may for instance beselected from acetic acid, propionic acid, butyric acid, isobutyricacid, n-valeric acid, trimethyl-acetic acid, caprioic acid and mixturesof one of these carboxylic acids with water, which in any event isproduced in the course of the reaction. The presently preferred solventis acetic acid and water. However, we do not preclude the possibility ofusing other solvents such as benzoic acid, e.g. a mixture of benzoicacid and water.

The water content used in the conventional production of terephthalicacid by liquid phase oxidation of paraxylene is typically such thatwater comprises between 3 and 10% by weight of the combinedsolvent-forming carboxylic acid/water supplied to the reaction zone. Afeature applicable to the various aspects of the invention disclosedherein is that the water content can be substantially greater than thatpresent in the the total feed to the reaction zone of a conventionalterephthalic acid production process; in the various aspects of theinvention disclosed herein the reaction medium composition at the timeof commencement of the reaction may contain water in an amount ofranging from about 3% up to about 30%, e.g. 12%, or greater (e.g. from10% up to 30%) by weight. The increased water content becomes feasiblebecause of the relatively high solvent to precursor ratio employed inthe process of the invention in order to ensure that substantially allof the terephthalic acid produced remains in solution during thereaction. For example, where the precursor comprises paraxylene, thesolubility of paraxylene in an acetic acid/water mixture falls sharplywith increasing water content and this imposes constraints on the amountof water that can be present in the reaction medium in conventionalterephthalic acid production since the solvent/paraxylene ratios arelow, typically between 4:1 and 7:1. Thus, in the process of the presentinvention, increased water content may be tolerable.

The solvent/precursor ratio is usually at least 30:1, e.g. at least50:1. In practice, it may be substantially greater than 50:1, forinstance up to 200:1, e.g up to 150:1. Where we refer to thesolvent/precursor ratio it is to be understood that the water componentpresent in the reaction medium forms part of the solvent and is to beincluded as such in determining the solvent/precursor ratio.

According to a third aspect of the present invention, which may be butis not necessarily used in conjunction with the first and/or secondaspect of the invention, there is provided a process for the productionof terephthalic acid by the liquid phase oxidation of a precursor ofterephthalic acid with oxygen in a reaction medium containing theprecursor and a solvent under conditions such that substantially all ofthe terephthalic acid produced in the oxidation reaction zone ismaintained in solution during the reaction, characterised in that thereaction is carried out by producing a flow of the reaction mediumthrough a reaction zone from an inlet region to an outlet region andestablishing a temperature profile along the direction of flow throughthe reaction zone such that the temperature of the reaction medium isgreater at the outlet region than at the inlet region.

The production of terephthalic acid by the oxidation of a precursorthereof is a highly exothermic reaction. Conventionally substantiallyisothermal conditions are maintained in the reactor through removing theheat of reaction by allowing solvent and water of reaction to vaporiseand removing the resulting vapour from the reactor. In the processaccording to this aspect of the present invention, the reaction iscarried out under non-isothermal conditions. Thus the heat of reactionneed not necessarily be removed or may only be removed to a lesserextent, with the consequence that the temperature within the reactionzone increases from the inlet region to the outlet region of thereaction zone. Typically the subsequent recovery of terephthalic acidfrom the reaction medium involves cooling of the latter to precipitatethe product and separation thereof from the resulting mother liquorwhich liquor may be recycled to the reaction zone. By allowing thetemperature of the reaction medium to increase on passage through thereaction zone, the temperature of recovered mother liquor may be suchthat little or no further heating of the recovered mother liquor isneeded prior to its reintroduction into the reaction zone. Thus, forinstance, the temperature of the mother liquor recovered followingprecipitation and separation of the terephthalic acid may differ fromthe temperature at the inlet of the reaction zone by no more than about30° C., more preferably no more than about 20° C.

It is to be understood that we do not exclude the possibility that, inthe course of the reaction, some precipitation of terephthalic acid fromsolution may occur for instance as a result of the temperature at one ormore locations along the direction of flow being insufficient tomaintain substantially complete dissolution; such precipitation can becompensated for by profiling the temperature along the direction of flowso that precipitated terephthalic acid redissolves at least in part.Also the temperature may be controlled so that any fine particles ofterephthalic acid present in any of the liquid phase components suppliedto the reaction zone undergo dissolution, at least in part, as thereaction medium progresses through the reaction zone. Such fineparticles may be introduced for example as a result of recycle of liquorseparated in the course of recovery of terephthalic acid from thereaction medium downstream of the reaction zone.

The temperature profile is typically established by allowing thetemperature of the reaction medium to increase, and/or by controllingthe temperature rise, due to the exotherm produced by the reaction.

The reaction zone may be formed by a single vessel or conduit or it maycomprise a series of sub-zones with each sub-zone formed by a separatevessel or conduit or by separate chambers within a single vessel.

The temperature profile from the inlet region to the outlet region mayincrease substantially continuously in the direction of flow or it maybe of a step-wise character. For instance, where the reaction zonecomprises a series of sub-zones there may be provision in at least onesuch sub-zone for removing or adding heat within that sub-zone in orderto establish a temperature profile which, preferably, is consistent withsubstantially all of the terephthalic acid being maintained in solutionthroughout the reaction zone.

There may be more than one reaction zone; for instance, two or morereaction zones in parallel each supplied with reactants and solvent and,if desired, the product streams from such multiple reaction zones may beunited to form a single product stream.

Where the heat of reaction is removed from the reaction zone (or one ormore sub-zone(s) thereof where applicable), it may be removed by heattransfer from the reaction medium to a heat sink across a heat exchangesurface, for instance heat exchange with a heat accepting fluid, and/orby introduction of a quenching liquid such as solvent and/or precursoror an immiscible liquid (as referred to in more detail hereinafter)which may be added in one or more stages along the direction of flow ofreaction medium through the reactor.

Where a heat accepting fluid is used, it is conveniently passed throughone or more flow passages having a wall or walls, the outer surfaces ofwhich are exposed to the reaction medium within the reaction zone. Forinstance, the heat accepting fluid may be circulated through a coiledtube or tubes immersed in the reaction medium. Alternatively, thereactor may be designed in a manner akin to a tube in shell heatexchanger with the reactants and solvent being passed through the shelland the heat accepting fluid being passed through the tubes internallyof the shell.

However, we do not exclude the possibility of effecting the thermaltransfer in other ways, for instance by passing the heat accepting fluidthrough a jacket arrangement at least partly surrounding the reactionzone. For example, the tube in shell design referred to above may besuch that the reactants and solvent flow through the tubes while theheat accepting fluid flows through the shell.

The heat accepting fluid may traverse the reaction zone incountercurrent and/or co-current relation with the reaction mediumflowing through the reaction zone. Conveniently the passage or passagesconducting the heat accepting fluid are arranged to extend internally ofthe reactor.

Advantageously the heat accepting fluid following heat exchange with thereaction medium is processed to recover thermal, mechanical and/orelectrical energy. The power recovered may in part be employed topressurise air or oxygen to be supplied as oxidant to the process, e.g.by driving a compressor suitable for this purpose. For example, heattransferred to the heat accepting fluid may be converted to mechanicalor electrical energy in a power recovery system. One approach is to usethe heat accepting fluid to raise steam which can then be superheatedand supplied to a steam turbine to recover power.

The heat accepting fluid may be preheated prior to traversing thereaction zone and such preheating may be effected by heat exchange withthe product stream resulting from the oxidation reaction.

Conveniently the heat accepting fluid comprises water or oil, e.g. amineral oil. Alternatively, the heat accepting fluid may comprise thereaction medium or one of the components thereof (i.e. solvent and/orprecursor). For instance, the exotherm generated during the reaction mayin part be used to preheat incoming reaction medium and in part used toraise the temperature of the reaction medium as it passes through thereaction zone so that substantially all of the terephthalic acid ismaintained in solution.

The initial temperature at the commencement of reaction will need to besufficiently high to ensure that the reaction is initiated but not sohigh that the temperature rise during the reaction leads to atemperature which results in excessive burning of solvent and aromatics.A typical inlet temperature range would be within the range 80 to 200°C., preferably between 120 and 180° C., e.g. between 140 and 170° C. Thetemperature of the product stream emerging from the reaction zone willbe in excess of the inlet temperature and may be between 180 and 250°C., preferably between 180 and 230° C. (e.g. 190 to 220° C.).

After traversing the reaction zone, substantially all of theterephthalic acid is in solution. The solution may also containcatalyst, and relatively small quantities of intermediates (e.g.p-toluic acid and 4-CBA) and by-products such as colour bodies andtrimellitic acid. The desired product, terephthalic acid, may beprecipitated, for instance by causing or allowing it to crystallise fromthe solution in one or more stages, followed by processing of theresulting slurry by solids-liquid separation in one or more stages.

Before the precipitation process, e.g. crystallisation, is implementedand while substantially all of the terephthalic acid and othercomponents are still in solution, the reaction medium may be treated soas to remove certain components. For example, the reaction medium may betreated to remove catalyst metal ions by an ion exchange techniquesusing for instance a cationic exchange resin or by electrodialysistechniques involving ion exchange membranes.

Because the product/mother liquor slurry resulting from precipitationand separation stages will be relatively thin in view of the relativelyhigh solvent:precursor ratios employed in the process of the presentinvention, preferably concentration of the product is effected upstreamof the solids-liquid separation. Concentration of the slurry may beeffected downstream of the crystallisation process using for instanceone or more hydrocyclone separators or it may be effected in the courseof the crystallisation process by using integratedcrystalliser/concentrating apparatus.

The solvent-based mother liquor (which may but need not necessarilycontain dissolved catalyst components) recovered following thesolids-liquid separation is preferably recycled to the oxidationreaction zone.

The recovery of the terephthalic acid may be effected by conventionalcrystallisation techniques involving reduction in pressure of thereaction medium. However, such pressure reduction gives rise to a needto repressurise the mother liquor to be recycled to the reaction system.

In another aspect of the present invention, which may be but is notnecessarily used in conjunction with the previously mentioned, there isprovided a process for the production of terephthalic acid by the liquidphase oxidation of a precursor of terephthalic acid with oxygen in areaction medium containing the precursor and a solvent under conditionssuch that substantially all of the terephthalic acid produced in theoxidation reaction zone is maintained in solution during the reaction,characterised in that the terephthalic acid is recovered from thereaction medium by precipitation in such a way as to avoid substantialdepressurisation of the reaction medium thereby allowing mother liquorto be recovered at a pressure which is substantially the same as thereactor operating pressure or a pressure close thereto, i.e. about 5bara or less, preferably about 2 bara or less, lower than the operatingpressure of the reactor system.

In each of the above-mentioned and following aspects of the inventiondisclosed herein, prior to re-introduction into the oxidation reactionzone, the mother liquor may be heated by heat exchange with the reactionmedium after the latter has emerged from the reaction zone and/or whilethe latter is traversing the reaction zone, thereby cooling the reactionmedium.

Usually the resulting precipitate will contain no more than 5000 ppm byweight of 4-CBA. Preferably the terephthalic acid precipitated from thereaction medium contains no more than 3000 ppm, more preferably no morethan about 1000 ppm and most preferably no more than about 500 ppm (e.g.20 to 300 ppm), by weight of 4-CBA.

Cooling of the reaction medium, with consequent precipitation of theterephthalic acid, is advantageously carried out, preferably undersuperatmospheric conditions, in such a way that the temperature of theresulting slurry undergoing solids-liquid separation is within the rangeof 120 to 180° C., more preferably about 130 to about 175° C. and mostpreferably about 140 to about 170° C. Although carrying out thesolids-liquid separation at such a high temperature results in asubstantial proportion of the terephthalic acid remaining in solution,it has been found that the levels of the major impurities in therecovered product reduce as temperature initially falls from thetemperature at which the reaction medium is withdrawn from the reactionzone and then increase as the temperature falls further.

According to a further aspect of the invention, which may be but is notnecessarily used in conjunction with other aspects of the inventiondisclosed herein, there is provided a process for the production ofterephthalic acid by the liquid phase oxidation of a precursor ofterephthalic acid with oxygen in a reaction medium containing theprecursor and a solvent under conditions such that substantially all ofthe terephthalic acid produced in the oxidation reaction zone ismaintained in solution during the reaction, the reaction mediumthereafter being cooled to precipitate terephthalic acid which isrecovered by solids-liquid separation, characterised in that thesolids-liquid separation is carried out at a temperature within therange of about 120 to about 180° C., more preferably about 130 to about175° C. and most preferably about 140 to about 170° C.

According to yet another aspect of the present invention, which may bebut is not necessarily used in conjunction with other aspects of theinvention disclosed herein, there is provided a process for theproduction of terephthalic acid by the liquid phase oxidation of aprecursor thereof in a solvent, the liquid phase oxidation being carriedout in a reaction zone in such a way that substantially all of theterephthalic acid produced is maintained in solution in the reactionmedium during the reaction, characterised in that the oxidant isintroduced into the reaction zone at two or more locations spaced apartin the direction of flow of the reaction medium from an inlet region ofthe reaction zone to an outlet region thereof.

This aspect of the invention is particularly applicable to the casewhere the reaction zone is formed, at least in part, by a plug flowreactor and is particularly beneficial where the oxidant is in the formof substantially pure oxygen or an oxygen enriched gas.

Such locations are conveniently so positioned relative to the bulk flowof solvent and reactants through the oxidation zone that oxidant isintroduced to the reaction at an initial location and at least onefurther location downstream of said initial location. The oxidant may beintroduced substantially continuously over a length of the reactionmedium flow path through the reaction zone and/or sub-zone(s); forexample, the oxidant may be introduced by means of a perforated pipeimmersed in the reaction medium and extending in the direction of flow,the number, spacing and distribution of the perforations being such thatthe oxidant is introduced at substantially all points along said lengthof the reaction zone and/or sub-zone(s).

The oxidant in each of the foregoing aspects of the invention isconveniently molecular oxygen, e.g. substantially pure oxygen, air orother oxygen containing gas (i.e. gas containing oxygen as the major orminor constituent thereof), or oxygen dissolved in liquid. The use ofsubstantially pure oxygen as the oxidant has the benefits of avoidinggas voidage and disruption of plug flow profile while affording highoxygen mass transfer rates required for intensified reaction at modestoperating pressures.

The oxygen may be combined with a diluent gas, such as carbon dioxide,which is more soluble in the solvent than nitrogen. The diluent gas maybe derived for instance from the vent gas produced during the oxidationreaction. Where the diluent gas is derived from the vent gas, the ventgas will preferably have been treated, e.g. by high temperaturecatalytic combustion, to convert any methyl bromide present to HBr andBr₂ and may be recycled, at least in part, without removing its HBrcontent since HBr can be employed as a catalyst component in theoxidation reaction. For example, following treatment to convert MeBr toHBr and Br₂, part of the vent gas may be diverted for dilution of theoxygen supply to the reaction while the remainder may be processedfurther, e.g. for disposal or use as a fluidising medium for conveyingpurposes. The diverted portion of the treated off gas may be cooled (forinstance, by heat exchange with the vent gas upstream of the MeBrconversion step) and recompressed (before or after admixture with theoxygen supply) sufficiently to allow it to be reintroduced in theoxidation reaction. The processing of the remaining vent gas maycomprise supply to a power recovery system such as an expander andscrubbing (upstream and/or downstream of the expander) to remove anyresidual pollutants such as HBr and Br₂.

Instead of molecular oxygen, the oxidant may comprise atomic oxygen suchas a compound, e.g. a liquid phase compound at room temperature,containing one or more oxygen atoms per molecule. One such compound forexample is hydrogen peroxide.

Besides the solvent:precursor ratio, various other parameters such astemperature and water content also need to be taken into account inorder to ensure that substantially all of the terephthalic acid producedis maintained in solution during the reaction.

The elevated pressure conditions under which the reaction is carried outwill normally be selected such that the reaction medium is maintained inthe liquid phase during the reaction (non-boiling conditions). Usuallythe reaction will be carried out at a pressure in the range of 10 barato 100 bara, typically 20 bara to 80 bara, depending on the nature ofthe oxidant; for instance if the reaction is carried out using dissolvedoxygen, the pressure is typically about 60 to about 80 bara wheresubstantially pure oxygen is employed but may be greater, e.g. above 100bara where the oxygen and a diluent are dissolved in the reactionmedium.

According to a further aspect of the invention, which may be but is notnecessarily used in conjunction with other aspects of the inventiondisclosed herein, there is provided a process for the production ofterephthalic acid by the liquid phase oxidation of a precursor ofterephthalic acid with oxygen in a reaction medium containing theprecursor and a solvent under conditions such that substantially all ofthe terephthalic acid produced in the oxidation reaction zone ismaintained in solution during the reaction, characterised in that thetotal oxidation reaction volume A, in m³, associated with the reactionzone, the 4-CBA content B of the recovered terephthalic acid in ppm w/w,and the amount of terephthalic acid C recovered from the oxidationreaction, in te/hr, are related by the formula:

(A*B)/C<4,000.

The term “total oxidation reaction volume” is to be understood tocomprise the total volume of the reactor vessel or vessels (in paralleland/or series) forming the reaction zone, including any vapour headspace provided in such vessel or vessels, e.g. for liquid/vapourdisengagement.

Preferably the the relationship is such that:

(A*B)/C<3,000.

Usually, in the context of this aspect of the invention, the 4-CBAcontent of the recovered terephthalic acid will be no greater than about5,000 ppm w/w and preferably is no greater than about 3,000 ppm w/w,more preferably no greater than about 1,000 ppm w/w, and may be lowerthan about 500 ppm w/w, e.g. in the range of about 20 to about 300 ppmw/w. Also the rate of production of terephthalic acid will typically bein excess of 20 te/hr.

This aspect of the invention may be implemented by operating theterephthalic acid process in accordance with the various aspects of theinvention; for instance, by carrying out the production process in anoxygen-fed continuous plug flow (or quasi-plug flow) regime innon-boiling conditions. For instance, a substantially pure oxygen-fedsingle plug flow reactor in accordance with the present inventionoperated to produce terephthalic acid product at a rate of 60 te/hr witha 4-CBA content of 250 ppm w/w can be implemented by a design having atotal oxidation reaction volume of less than 160 m³ and, in this event,(A*B)/C<1,000. In contrast, a conventionally designed oxidation reactorcurrently in operation and designed to produce 60 te/hr crudeterephthalic acid with a 4-CBA content of about 2500 ppm w/w requires atotal oxidation reaction volume in excess of 400 m³ and therefore, inthis case, (A*B)/C>16,000, which clearly demonstrates the significantreduction in reactor volume that may be achieved by virtue of thepresent invention. In the case of a conventional single stage oxidationprocess using multiple reactors operating with high catalystconcentration and high temperature (and consequent high levels of aceticacid burning) to produce fibre grade terephthalic acid with a 4-CBAcontent of 500 ppm and at a rate of 60 te/hr, the total oxidationreaction volume required is of the order of about 800 m³ which gives avalue greater than 6,600 for the relationship (A*B)/C.

In a related aspect of the invention, plant for the production ofterephthalic acid by the liquid phase oxidation of a precursor thereofand operationally designed to produce terephthalic acid having a 4-CBAcontent B of less than about 5,000 ppm w/w at a production rate C of atleast 20 te/hr, characterised in that the total oxidation reactionvolume A, in m³, of the vessel or vessels in which the oxidationreaction is carried out satisfies the following condition:

A<(4,000*C)/B.

The process in any one of the foregoing aspects of the invention willnormally be carried out in the presence of an oxidation catalyst. Whereemployed, the catalyst may be soluble in the reaction medium comprisingsolvent and the terephthalic acid precursor(s) or, alternatively, aheterogeneous catalyst may be used. The catalyst, whether homogeneous orheterogeneous, typically comprises one or more heavy metal compounds,eg. cobalt and/or manganese compounds, and may optionally include anoxidation promoter such as bromine or acetaldehyde. For instance, thecatalyst may take any of the forms that have been used in the liquidphase oxidation of terephthalic acid precursors such as terephthalicacid precursor(s) in aliphatic carboxylic acid solvent, eg. bromides,bromoalkanoates or alkanoates (usually C1-C4 alkanoates such asacetates) of cobalt and/or manganese. Compounds of other heavy metalssuch as vanadium, chromium, iron, molybdenum, a lanthanide such ascerium, zirconium, hafnium, and/or nickel may be used instead of cobaltand/or manganese. Advantageously, the catalyst system will includemanganese bromide (MnBr₂). The oxidation catalyst may alternatively oradditionally include one or more noble metals or compounds thereof, e.g.platinum and/or palladium or compounds thereof, for example in highlydivided form or in the form of a metal sponge. The oxidation promoterwhere employed may be in the form of elemental bromine, ionic bromide(eg. HBr, NaBr, KBr, NH₄Br) and/or organic bromide (eg. bromobenzenes,benzyl-bromide, mono- and di-bromoacetic acid, bromoacetyl bromide,tetrabromoethane, ethylene-di-bromide, etc.). Alternatively theoxidation promoter may comprise a ketone, such as methylethyl ketone, oraldehyde, such as acetaldehyde.

Where the catalyst is in heterogeneous form, it may be suitably locatedwithin the reaction zone so as to secure contact between thecontinuously flowing reaction medium and the catalyst. In this event,the catalyst may be suitably supported and/or constrained within thereaction zone to secure such contact without unduly constricting theflow cross-section. For instance, the heterogeneous catalyst may becoated on or otherwise applied to, or embodied in, static elements (eg.elements forming an openwork structure) positioned within the reactionzone so that the reaction medium flows over the same. Such staticelements may additionally serve to enhance mixing of the reactants asthey pass through the reaction zone. Alternatively the catalyst may bein the form of mobile pellets, particles, finely divided form, metalsponge form or the like with means being provided if necessary toconfine the same to the reaction zone so that, in operation, thecatalyst pellets etc become suspended or immersed in the reaction mediumflowing through the reaction zone. The use of a heterogeneous catalystin any of these ways confers the advantage of being able to confine thecatalysis effect to a well-defined zone so that, once the reactionmedium has traversed the zone, further oxidation takes place at areduced rate or may be significantly suppressed. Also, provision forcatalyst recovery may be avoided.

The support for the oxidation catalyst can be less catalytically activeor even inert to the oxidation reaction. The support may be porous. Ingeneral, the catalyst support materials will be substantially corrosionresistant and substantially oxidation resistant under the conditionsprevailing. Thus, depending on the prevailing conditions, the catalystsupport material can be selected from for example titania, silica,alumina, silica alumina, alpha alumina, gamma alumina, delta alumina,and eta alumina, mullite, spinel, andzirconia. Supports comprising alphaalumina, gamma alumina, silica, or silica alumina are preferred.

The support component of the oxidation catalyst may be pure or acomposite of materials, the latter being employed for example to impartdesired chemical or physical characteristics to the catalyst. Forinstance, the oxidation catalyst may comprise a substrate with highattrition resistance and a substrate coating having high surface area.Conventional impregnation techniques may be used to fabricate the same.Materials for use as the substrate will generally be substantiallycorrosion resistant and substantially oxidation resistant under theconditions prevailing. Thus, depending on the prevailing conditions, thesubstrate material can be selected from alpha alumina, mullite andspinel. Materials for use as a composite substrate coating are silica,alumina, titania, zirconia, alpha alumina, gamma alumina, delta aluminaand eta alumina.

The invention will now be described by way of example only withreference to the accompanying drawings illustrating application of theprocesses according to various aspects of the invention to theproduction of terephthalic acid. In the drawings:

FIG. 1 is a block diagram showing the overall oxidation process;

FIG. 1A is a view showing one method of combining the various feeds toform the reaction medium;

FIG. 2 is a flow sheet illustrating one form of oxidation reactor schemethat may be used in the process of the present invention employingsubstantially pure oxygen or oxygen enriched gas as the oxidant;

FIG. 3 illustrates a modification of the oxidation reactor shown in FIG.2;

FIG. 4 is an alternative scheme in which the reaction is carried outusing continuous stirred tank reactors;

FIG. 5 is a flow sheet of one embodiment for use in the crystallisationand recovery of terephthalic acid;

FIG. 6 is a flow sheet illustrating use of a rectifier to recoversolvent and water from flash vapours produced in the crystallisationprocess;

FIG. 7 is a flow sheet illustrating a product recovery scheme in whichcrystallisation and concentration are carried out in the same vessel;

FIG. 8 illustrates one form of plug flow reactor system with provisionfor heat removal; and

FIG. 9 is a schematic view of apparatus used in experimental work toproduce the Examples reported herein.

Referring to FIG. 1, terephthalic acid is produced in a reactor system10 by the liquid phase oxidation of a precursor thereof, e.g.paraxylene, in a solvent such as acetic acid, the oxidation beingcarried out in the presence of a catalyst system. The reactor system 10may take various forms such as a single plug flow reactor, two or moreplug flow reactors arranged in series, a plug flow reactor incombination with one or more continuous stirred tank reactors, or two ormore continuous stirred tank reactors in series and arranged so as toapproximate plug flow. Some examples of the possible reactor systemconfigurations will be described below. Precursor, make-up solvent,make-up catalyst (e.g. comprising cobalt and manganese compoundstogether with bromine as an oxidation promoter) and recovered motherliquor and solvent components are mixed in mixer and preheater section12 to produce a reaction medium in which the solvent (fresh andrecovered) to precursor ratio in the mixture is substantially higherthan that used in the conventional oxidation of paraxylene terephthalicacid by liquid phase oxidation. At least part of the preheat (whereneeded) may be provided by the recovered mother liquor and solvent understeady state operating conditions of the system. The heat suppliedthrough the recovered mother liquor may be sufficient to eliminate theneed for an external heat source under steady state conditions although,in this event, an external source will still be needed on start-up.Typically the solvent:precursor ratio is of the order of about 70:1 (ona weight basis). The mixture is supplied via line 14 to the inlet regionof the reaction system. The temperature of the mixture supplied to thereaction system is typically at a temperature of about 150° C. and ispumped to a suitable pressure to ensure that boiling of reaction mediumduring the reaction is substantially prevented. Instead of mixing theliquid phase components in mixer 12, they may instead be preheated butkept separate until introduction into the reactor system (in which caseunit 12 may simply be a preheater) and supplied to the reactor system 10as two or more separate feeds 14 to be mixed at the inlet region of thereactor system 10 (see FIG. 1A for an example of such a scheme).

Oxygen is supplied via line 16. The oxygen supply may take various formsincluding substantially pure oxygen, air, oxygen enriched air, gascontaining oxygen and a diluent such as nitrogen or carbon dioxide etc.Although the oxygen supply is depicted by a single line 16 entering thereaction system, the method of supplying oxygen and the nature of theoxygen supply in terms of its concentration may vary as will becomeapparent from the more specific embodiments described below. Also whilethe source of oxygen is shown as being supplied to the reactor systemseparately from the solvent/paraxylene reaction medium, it is preferablypredissolved at least in part (see line 17) in the reaction medium orone or more components thereof (e.g. acetic acid and/or mother liquorrecycle) upstream of the reactor system or the mixer/preheater section12, irrespective of the oxygen source used but particularly wheresubstantially pure oxygen, or oxygen diluted with an inert gas, is used.

The solvent/precursor/catalyst reaction medium passes through thereactor system, preferably as a plug flow or plug flow approximation,from the inlet region to an outlet region at which a product stream iswithdrawn via line 18. The reaction is carried out in such a way thatsubstantially all of the terephthalic acid formed during passage of thesolvent/precursor mixture through the reaction is maintained insolution, thereby maintaining intermediates such as paratoluic acid and4-CBA in solution during the reaction and hence available for reaction.In this manner, it is possible to secure a product having a relativelylow 4-CBA content.

The product stream is passed via line 18 to a crystallisation section 20in which precipitation of the product, terephthalic acid, is effected toform a thin slurry of terephthalic acid in a mother liquor which mainlycomprises the solvent employed and some water, dissolved catalystcomponents, terephthalic acid, intermediates thereof and by-productsformed in the reaction. The crystallisation process involves reducingpressure and temperature and at the end of the process the slurrypressure may range from below atmospheric pressure to pressuressubstantially above atmospheric pressure, preferably the latter.

The temperature at which the crystallisation process is terminated maybe selected so that the mother liquor recovered subsequently from theslurry is at a suitable temperature such that, when mixed with make-upsolvent and precursor, the mixed stream has a predetermined temperaturecorresponding to the desired inlet region temperature of the reactionzone. Water is produced in the oxidation reaction; one method ofremoving at least part of the water of reaction is to use a pressurerectifier/distillation column in conjunction with the crystallisationsection; for example by supplying the flash vapour from at least one ofthe crystalliser vessels, at elevated pressure, directly to adistillation column for separating solvent (as bottoms product) from thewater (overheads product in the form of steam). The pressurised steamoverheads may then be used in a power recovery system by means of asteam condensing turbine. One example of such a scheme will be describedin more detail below with reference to FIG. 6.

Where the terephthalic acid is precipitated by a crystallisation processinvolving reduction in the pressure of the reactor product stream belowits saturated vapour pressure to initiate solvent removal by flashingand solvent cooling, following recovery of the terephthalic acid, atleast the bulk of the residual mother liquor is re-pressurised andrecycled to the reactor. In an alternative approach aimed at avoidinghaving to repressurise the mother liquor recycle stream, theterephthalic acid may be precipitated by cooling the reactor productstream without reducing its pressure. Heat is then removed via a heatexchange surface and used in, for example, steam raising or processheating etc. In such an arrangement, on cooling the terephthalic acidwill tend to foul the heat exchanger surface, reducing itseffectiveness. This fouling can be managed by employing a surfacescraped heat exchanger device.

Another alternative with the aim of reducing the extent to whichre-pressurisation of the mother liquor recycle stream is necessaryinvolves precipitating the terephthalic acid from the reaction medium bysolvent removal, without cooling the solvent. Solvent removal can beeffected by forcing it through a semi-permeable membrane (permeable toacetic acid and water and optionally to catalyst and reactionimpurities, but impermeable to terephthalic acid). On solvent removal,terephthalic acid precipitation is initiated and fouling of the membranepores, which would otherwise reduce membrane effectiveness, can becountered by, for example, design for high shear across the membranesurface and/or staging of membranes in series with intermediate vesselsin which the bulk of the crystallisation occurs. Because pressure dropthrough the membrane system is not substantial, the pumping cost foreffecting recycle of mother liquor can be reduced.

Following the crystallisation process, the slurry strength will besignificantly lower than is the case in the conventional production ofterephthalic acid—i.e. because of the high solvent:precursor ratioemployed. Desirably therefore, before carrying out solids-liquidseparation, the slurry is concentrated. This can be effected downstreamof the crystallisation section in a concentration section 22 which may,for example, comprise one or more hydrocyclone stages producing athickened underflow stream comprising the major part of the terephthalicacid crystal mass in the slurry and mother liquor and an overflow streamcomprising mother liquor in which terephthalic acid fines may besuspended. Concentration of the slurry by means of one or morehydrocyclones is particularly expedient in view of the relatively lowcost of such devices; however, other devices may be used instead such asone or more centrifuges (nozzle, decanter etc), a filter or filters(including cross-flow microfiltration), or gravityclarification/thickening devices either separate from or incorporated inthe crystalliser (as described hereinafter).

The overflow stream from the concentration section 22 is routed vialines 24 and 26 for recycle to the mixer/preheater 12. The concentrateis supplied to a solids-liquid separation section 28 in which theterephthalic acid crystals are separated from the mother liquor, thesolids-liquid separation being carried out using for example one or morefiltration devices operating under superatmospheric, atmospheric orsub-atmospheric conditions, with or without washing facilities, such asdescribed in our prior published International Patent Applications Nos.WO 93/24440 and WO 94/17982 (the entire disclosures of which areincorporated herein by this reference). Thus, for example the integratedsolids separation and water washing apparatus may comprise a centrifuge,a belt filter unit, a rotary cylindrical filter unit operated with theslurry side under pressure, or a drum filter unit (e.g. a BHS-Festpressure filter drum formed with a plurality of slurry receiving cellsin which the mother liquor is displaced from filter cake by water underhydraulic pressure supplied to the cells). After filtering the slurry,the recovered terephthalic acid may be dried. If not already atatmospheric pressure, the filter cake of terephthalic acid may betransferred to a low pressure zone (e.g. atmospheric pressure) fordrying via a suitable pressure letdown device such as a lock hopperarrangement, a rotary valve, a ram-type pump, a screw feed device or aprogressive feed device such as a progressive cavity pump of the typeused to pump cold pastes of high solids contents.

The temperature of separation and the degree of washing required will bedependent on the levels of impurities generated in the reaction and therequired product specification. Although, in general, it will bedesirable to produce terephthalic acid which is sufficiently pure torender further purification unnecessary (e.g. by oxidation and/orhydrogenation of an aqueous solution of the terephthalic acid to convert4-CBA to terephthalic acid or to paratoluic acid, as the case may be),we do not exclude the possibility of carrying out such purification inthe process of the present invention. Following solids-liquidseparation, the product may be recovered via line 30 for drying and usein the downstream production of polyester by esterification with a diol(e.g. ethylene glycol) without necessarily requiring interveningchemical purification. Drying of the product may be carried out in forexample a rotary steam tube drier or a fluidised bed drier.

The mother liquor obtained as filtrate from the solids-liquid separationsection 28 is routed via lines 32 and 26 for recycle to themixer/preheater 12. The mother liquor may contain some solid phaseterephthalic acid in the form of fines. This fines content may at leastin part redissolve in the reaction medium as a result of preheating orwithin the reaction system; however, even if some of the fines contentremains undissolved as it passes through the reaction system, this willnormally be tolerable since the fines will tend to be relatively pure incontrast with the fines produced in the conventional process for theproduction of terephthalic acid. Although not illustrated in FIG. 1, therecovered mother liquor (lines 32 and 26) may be heated before return tothe mixer 12 using a heat exchanger-to effect heat transfer from theproduct stream on line 18 to the mother liquor recycle. Alternatively,it may be desirable to cool the mother liquor (e.g. where the motherliquor is used as a vehicle for pre-dissolving oxygen) in which case themother liquor may for example be brought into heat exchange relationwith the feed or one of the feeds 14 to the reactor system so as toeffect heating of the feedstream or feedstreams.

Part of the solids recovered downstream of the crystallisation processmay be recycled back to one or more of the crystallisers in order to“seed” the solution and nucleate and/or promote particle growth. Forinstance, part of the fines-containing overflow stream or the thickenedunderflow stream from the concentration section may be recycled to thecrystallisation process for this purpose.

The crystallisation process typically involves flashing off solvent andwater from the slurry, the water being produced as a reactionby-product. The resulting vapour and/or condensate is supplied via line35 to a solvent recovery section 34. Solvent recovered in the solventrecovery section is routed via lines 38 and 26 to the mixer/preheater 12while gases and other volatiles are passed to a vent treatment system 40via line 42 along with any volatiles and gases, including unreactedoxygen, recovered from the reaction system and/or crystalliser sectionvia line 44. Where the reactor system is operated under sufficientlyhigh pressure conditions to secure a single phase regime throughout thereaction, no vent gases are obtained from the reactor system; insteadthe gaseous components are removed by venting when they come out ofsolution in the crystallisation section. Water separated from thesolvent in solvent recovery section 34 is routed via line 46 to aneffluent treatment plant.

The method of introducing the oxygen into the reaction may vary. In apreferred embodiment of the invention the oxygen or oxygen-containinggas is introduced into the reaction medium in such a way thatsubstantially all of the oxygen or oxygen-containing gas is dissolved inthe reaction medium so that the reaction can be conducted under singlephase conditions with those components which, in a conventionalterephthalic acid production process, would otherwise be in the gasphase and solid phase being present in dissolved form in the liquidphase reaction medium. FIG. 1A illustrates one scheme for achievingthis. In this case, the reactor 10A is supplied with the followingliquid phase components:

A. paraxylene and acetic acid solvent (liquid phase);

B. make-up catalyst in acetic acid solvent (liquid phase);

C. mother liquor recovered from the process (liquid phase); and

D. oxygen or oxygen-containing gas (gas phase).

Feeds A And B, which are relatively small in volume compared with feedC, are pumped to system pressure and then initially mixed together inmixer M1 and preheated if necessary during, before or after mixing,thereby producing combined feed E. Feed E is fed into the inlet region10A of the reactor system. Oxygen in excess of the stoichiometric amountrequired for the reaction is added, via feed D, to the mother liquorfeed C which is under system pressure and, if necessary, after additionof the oxygen is preheated, and the resulting oxygen-containing liquid,feed F, is introduced into the inlet region 10A. The oxygen may forinstance be added as a single jet into the mother liquor recycle streamimmediately upstream of a static mixer M2. The static mixer M2 isdesigned to ensure that: by continuous mixing of the mother liquorstream, there are no localised high concentrations of oxygen insolution; maximum bubble size is controlled by preventing bubblecoalescence; and bubbles are distributed uniformly throughout the motherliquor recycle stream so as to minimise the time taken to dissolve allof the gas in the liquid stream. The inlet region 10A includes a mixingarrangement M3, e.g a static mixer, for ensuring thorough mixing of thefeeds E and F thereby forming the reaction medium.

The system operating pressure, i.e. the pressure to which the feeds Eand F are pressurised, is selected so that all of the oxygen oroxygen-containing gas introduced enters into solution into the liquidphase reaction medium while ensuring that boiling of the reaction mediumis prevented. Where pure oxygen is employed as the oxygen source, thesystem operating pressure may typically be in excess of about 60 baraand will be correspondingly increased where a diluent is present. Forinstance, where the oxygen is supplied in the form of a gas containing80% oxygen and 20% nitrogen, the system operating pressure will betypically in excess of about 75 bara. The temperature of the feeds E andF will be such that, when combined, a desired reaction mediumtemperature (e.g. 150° C.) is secured at the inlet of the reactor system10 consistent with initiating the oxidation reaction.

The acetic acid/paraxylene ratio will be determined by solventintroduced via feeds A and B and also by acetic acid recycled via themother liquor feed C. This ratio will be such that substantially all ofthe terephthalic acid produced in the ensuing reaction is maintained insolution throughout the reaction zone, taking into account the fact thatmore and more terephthalic acid is produced as the reaction mediumprogresses towards the outlet region 10B of the reactor system and alsothat the temperature of the reaction medium increases since the reactorsystem is operated non-isothermally resulting in a temperature increasefrom the inlet region 10A to the outlet region 10B. The temperatureprofile produced may be tailored and one such means for controlling thetemperature profile is to introduce feed F into the reactor system 10 instages rather than in one-shot at the inlet region 10A. Thus, asillustrated in phantom outline in FIG. 1A, the feed F may be split intoseparate feeds F, F1, F2 . . . with feed F being introduced at the inletregion 10A and the remaining feeds F1, F2 . . . being injected as quenchfeeds into the reaction medium at different points along the path offlow of the reaction medium through the reactor system. At eachinjection point, suitable mixing arrangements M4, M5 will be provided toensure thorough mixing of the injected liquid with the remainingreaction medium. Where the reactor system comprises two or more separatereactor vessels, the feeds F1, F2 . . . may be conveniently injectedinto the reaction medium at the transitions between successive reactorvessels.

In practice, the mother liquor feed C is the most suitable vehicle forthe introduction of the oxygen into the reaction as it will normallyconstitute the bulk of the reaction medium during steady state operatingconditions; however, we do not exclude the possibility of oxygen beingintroduced by way of one or more of the other feeds to the reactorsystem either instead of the mother liquor feed or in addition thereto.

Referring now to FIG. 2, this illustrates one form of the processdescribed generally with reference to FIG. 1 in which the reactor systemis in the form of an adiabatic plug flow reactor 60 having an inletregion to which the reactants are supplied for mixing, the liquid phasereaction medium being produced by combining a mixture of paraxylene,fresh solvent and make-up catalyst supplied via line 62 frommixer/preheater 66, and recycled mother liquor feed on line 64. Theoxidant (in this case substantially pure oxygen) is supplied viacompressor 68 and line 70 and is combined with the mother liquor recyclefeed 64. The reactants are combined with suitable intensive mixing so asto produce a single phase with the oxygen dissolved in the liquid phase.Typically the liquid phase reaction medium is pumped to a pressure ofthe order of 60 bara and the oxygen is compressed to a pressure inexcess of 60 bara to facilitate its introduction into the liquid phase.As in the case of the embodiment of FIG. 1, the reaction mediumcomprises a mixture of paraxylene (pX), solvent-based mother liquor andsolvent recovered downstream of the reactor system, solvent make-up andcatalyst make-up. Typically the reaction medium/oxidant mixture suppliedto the inlet 62 of the reactor system is at a temperature of the orderof 150° C. Single phase operation of the reactor system can lead toadvantages in terms of reduction/suppression of methyl acetate formationwhere acetic acid is employed as solvent, and lower levels of solventburn can be achieved compared with conventional reactors. Moreover theuse of oxygen rather than air results in a substantial reduction in thevolume of gas vented during the subsequent crystallisation process.

The entire oxygen supply for the reaction may be introduced at the inletregion of the reaction system 60 in which case, to achieve single phaseconditions in which the oxygen is dissolved in the liquid phase, thiswould require operation of the reactor at high pressure (e.g. of theorder of 60 to 100 bara). If desired, the reactor pressure can bereduced significantly by distributing the oxygen supply along thedirection of flow through the reactor 60. Thus, as shown in FIG. 2, partof the oxygen is supplied by line 70 and the remainder is injected (withsuitable intensive mixing to secure single phase conditions) at a seriesof locations along the length of the reactor 60 via N injection points70A, 70B and 70C (where N is equal to one or more). In a variation asillustrated in FIG. 3, the distribution of the oxidant may be madesubstantially continuous by introducing it via a perforated conduit orthe like extending internally and lengthwise of the reactor 60. Byintroducing the oxidant progressively as illustrated in FIG. 2 or FIG.3, the amount of oxidant supplied to the reaction at a particularlocation can be tailored to the oxygen requirement at that location andto ensure that there are no oxygen starved zones within the reactor.Moreover, the reactor pressure can be reduced provided that it is keptat a level sufficient to prevent boiling of the reaction medium duringthe reaction.

Where the solvent:paraxylene ratio of the mixture supplied to the inletof the reactor 60 is for example of the order of about 60:1, theadiabatic temperature rise in the course of the reaction is about 70° C.The reaction exotherm may be removed by allowing the temperature to risefrom an inlet temperature of for example 150° C. at the reactor inlet 62to about 220° C. at the reactor outlet 72 without the need to removeheat, e.g. by indirect heat transfer or quenching by introduction ofcooler liquid (i.e. reaction medium or heat carrier). However, it willbe understood that the invention, including the embodiment of FIG. 2, isnot limited to removal of the exotherm solely by allowing thetemperature to rise from the inlet to the outlet of the reactor, andthat any suitable heat removal method may be employed to remove the heatof reaction at least in part.

The product stream withdrawn from the exit 72 of the plug flow reactoris processed in a product and solvent recovery section 74 correspondinggenerally to the plant sections 20, 22, 28 and 34 described withreference to FIG. 1 with recovered product, recovered water andrecovered mother liquor, mother liquor purge being depicted byreferences 76, 78, 80 and 82 respectively. Because in this embodiment,substantially pure oxygen is employed as the oxidant, nitrogen or otherinert gas is supplied to the crystallisation process via line 84 toensure that the vent gas recovered via line 86 is not flammable. Thenitrogen will usually be supplied to the headspace of at least the first(highest temperature, pressure) crystallisation vessel in thecrystallisation train. The vent gas stream 86 is passed through heatexchanger 88 where it is preheated by hot treated vent gas supplied vialine 90 before being heated in furnace 92 and then subjected tocatalytic combustion in unit 94 in order to destroy pollutants such ascarbon monoxide and organics which are convertible to CO₂ and water. Thetreated gas, following passage via line 90 through heat exchanger 88 canbe discharged, if necessary after scrubbing with water or alkali toremove any remaining pollutants such as bromine and HBr arising fromcatalytic combustion of any methyl bromide present in the vent gasstream 86.

In a variation of the embodiment of FIG. 2, there may be two or moreplug flow reactors connected in series, optionally with multipleinjection of oxidant into one or more of the series connected reactors.For example, product-containing stream from the first reactor may bepassed directly to the next reactor (and so on where more than two plugflow reactors are provided). A quenching medium may be introduced intoone or more of the reactors downstream of the first in order to controlthe temperature profile of the reactor system as a whole, consistentwith maintaining substantially all of the terephthalic acid formed insolution. The quenching medium will usually comprise the solventemployed in the reaction (e.g. mother liquor as described with referenceto FIG. 1A), optionally with oxygen predissolved therein. In anothervariation, the heat controlling medium introduced into one or more ofthe reactors downstream of the first reactor may, depending on thetemperature profile to be established through the reactor system, serveto heat rather than cool the product-containing stream undergoingtransfer from the preceding reactor.

In the foregoing description, the catalyst employed is dissolved in thesolvent medium supplied to the oxidation reaction; however, as discussedpreviously, a heterogeneous catalyst may be employed. Preferably thecatalyst system includes zirconium. For instance, supplementing thecobalt by 15% by weight of zirconium has been found to produce a markedreduction in 4-CBA content compared with identical conditions (residencetime, feed composition and temperature) but without such supplement—i.e.about 100 ppm 4-CBA with zirconium substitution in the catalyst systemcompared with 250 ppm 4-CBA without zirconium supplement.

In the embodiment of FIG. 2, the oxidation reaction is carried out inone ore more plug flow reactors. FIG. 4 illustrates an alternativeapproach in which the reaction zone is formed by a series of continuousstirred tank reactors 170A, 170B and 170C, optionally in combinationwith a small plug flow reactor 172. In this embodiment, plug flow-likeoperation is approximated by the use of multiple CSTRs—the greater thenumber of CSTRs employed, the closer the reactor system approachesoperation in a plug flow regime and hence a more favourable burn versus4-CBA relationship. The reaction medium from mixer/preheater 174 has thecomposition described with reference to the embodiment of FIG. 2 and ispressurised by pumping for supply to the inlet 176 of the reactionsystem, namely the inlet of the first CSTR 170A of the series, thereaction medium being pressurised to a pressure allowing sufficientmargin to avoid any significant boiling thereof in each CSTR, e.g. apressure of about 25 bara. The effluent from each CSTR is supplied tothe next via lines 178 and 180 and to the plug flow reactor 172 (wherepresent) via line 182. Air or oxygen in a diluent gas such as CO₂ iscompressed by compressor 184 to a pressure of about 32 bara and issupplied to each CSTR via line 186 and also to the plug flow reactor 172when present. In a modification, substantially pure oxygen may be usedas the oxidant supplied to the CSTRs and/or the plug flow reactorprovided that suitable precautions are taken with regard to the hazardsassociated with the presence of high levels of oxygen in the system. Forinstance, the CSTRs may be supplied with air while the plug flow reactormay be supplied with substantially pure oxygen or an enriched oxygensource such as oxygen combined with a diluent gas such as CO₂ ornitrogen. Regardless of the form in which the oxygen is supplied, theair/oxygen flow to each CSTR will desirably be controlled on exit oxygento that CSTR to ensure that it does not become starved of oxygen.

In the reaction zone, the temperature of the reaction medium isengineered in such a way that substantially all of the terephthalic acidproduced by oxidation of the paraxylene on that pass through thereaction zone is maintained in solution. In this context, it will beappreciated that there may be some terephthalic acid present in thesolid phase in the form of undissolved fines from the mother liquorrecycle as mentioned previously. Thus, in one scenario where thereaction medium supplied to the inlet region of the reaction zone is ata temperature of the order of 150° C., the first CSTR 170A may operateat a temperature of about 180° C., the second at a temperature of about200° C. and the third at a temperature of about 210° C. therebydeveloping a temperature profile consistent with maintainingsubstantially all of the terephthalic acid formed in solution. In otherwords, as the reaction progresses and more terephthalic acid forms, thereaction medium is passed to the next CSTR where the temperature issufficient to maintain in solution the terephthalic acid already formed,and that which will be formed in the CSTR concerned. The same pressuremay prevail within each CSTR but we do not exclude the possibility ofoperating the reactor system so that the pressure is increased from oneCSTR to the next.

Depleted air is removed from each CSTR 170A, B, C in an overheads streamwhich passes through a respective condenser 190A, B, C. These condensersare not for condensation of bulk boiloff (as in the conventional design)since the CSTRs are operated in a non-boiling regime; instead thecondensers serve to “knock back” some of the solvent that may beentrained as vapour or droplets in the depleted air stream. Theresidence time of the reaction medium in each CSTR will be relativelyshort compared with the residence time, typically 30 minutes or more, ina conventional terephthalic acid-producing CSTR. Typically the residencetime in each CSTR 170A, B, C may be of the order of several minutes, e.gabout 1 to 2 minutes per CSTR.

The product stream withdrawn from the final CSTR 170C may be passeddirectly to the product/solvent recovery section 192 but where low 4-CBAterephthalic acid product is desired without excessive burn of thesolvent, a small plug flow reactor 172 can be employed to further reducethe 4-CBA and other impurity content of the product exiting the finalCSTR 170C. The plug flow reactor employed for this purpose need only berelatively small—for example where the CSTR's may each have a volumetriccapacity of the order of 100 m³, the plug flow reactor 172 may have acapacity of the order of 10 to 20 m³. Thus, for example, the major part(e.g. at least 75%) of the terephthalic acid obtained in the oxidationreaction may be produced in the CSTRs with the remainder being producedin the plug flow reactor.

The product/solvent recovery section 192 will be generally similar tothat described with reference to FIGS. 1 and 2. Lines 194, 196, 198 and200 respectively depict the terephthalic acid product recovered,recovered water, mother liquor recycle to the mixer/preheater 174 andmother liquor purge. The depleted air stream from the condensers 190A,B, C will be at high pressure, e.g. 25 to 30 bara, and will containinter alia residual oxygen, carbon monoxide, carbon dioxide, asubstantial amount of nitrogen, solvent and other organics such asmethyl bromide and methyl acetate. This vent gas stream 202 is furthercooled in condenser 204 and the condensate, primarily solvent, is fedvia line 206, to the solvent recovery section. The cooled vent gasstream is then contacted with scrubbing liquor in high pressure absorber209 to remove further organics. The scrubbed vent gas is preheated infuel-fired preheater 210, subjected to catalytic combustion in unit 212and passed through expander 214 to recover power. Thenitrogen-containing gas stream 216 recovered from the expander 214 maythen be processed further, e.g. by scrubbing with water or alkali, toremove any residual pollutants such as bromine and/or HBr before beingdischarged and/or used for other duties, e.g. inerting duties, in theproduction plant. The vent gas treatment process may be generally asdescribed in published International Patent Application No. WO 96/39595.

In a modification of the embodiment of FIG. 4, where the plug flowreactor is present, a single CSTR may be used instead of a series asillustrated. The arrangement may for instance be such that the majorpart of the oxidation reaction is conducted in the CSTR, e.g. of theterephthalic acid product obtained from the overall reaction, at least75% is produced in the CSTR and the remainder produced in the plug flowreactor. In another modification, the or each CSTR may be supplied withsubstantially pure oxygen or an enriched oxygen gas (i.e. an oxygenconcentration in excess of that present in air—e.g. 23 to 100%) designedin the manner disclosed in U.S. Pat. No. 5,371,283 with means forestablishing a quiescent zone within the reactor in such a way that theoxygen bubbles are confined to the recirculating body of liquid and aresuppressed from entering the reactor headspace. As disclosed in U.S.Pat. No. 5,371,283 (the entire disclosure of which is incorporatedherein by this reference), this may be effected by means of a bafflelocated in the region of the interface between the gas and liquid phasesand/or by flooding the headspace within the reactor with an inert gassuch as nitrogen. In this manner, the oxygen content of the vent gas maybe monitored relatively easily in order to avoid flammability hazards inthe reactor headspace.

In those embodiments described above where one or more plug flowreactors are used, the reactor is shown orientated with its longitudinalaxis extending horizontally. However, it will be appreciated that thisis not essential and that the reactor may for instance be orientated sothat the flow of liquid takes place in a generally vertical direction.

One form of product recovery section as depicted in FIGS. 1, 2 and 4 isillustrated in FIG. 5 to which reference is now made. The product streamon line 220 from the reactor system enters a first stirred crystalliservessel 222A in which it is flashed to lower pressure and temperatureresulting in partial precipitation of the terephthalic acid content ofthe product stream and evolution of vapour comprising solvent and water.The product stream is next passed via line 224 to a second stirredcrystalliser vessel 222B in which it is flashed to even lower pressureand temperature with consequent further precipitation of terephthalicacid crystals and evolution of solvent and water vapour. Typically, fora product stream on line 220 at a temperature of the order of 210 to220° C., the product stream will be flashed to about 195° C. and 9 barain the first crystalliser and to about 151° C. and 3 bara in the secondcrystalliser. Although in FIG. 5 two stages of crystallisation areillustrated, it will be appreciated that there may be more than twostages or even a single stage of crystallisation.

The product stream comprising precipitated terephthalic acid crystals insolvent-based mother liquor is in the form of a relatively thinslurry—e.g. containing about 3.5 wt % solids. This slurry is fed viapump 226 to one or more hydrocyclones 228 (only one is illustrated inFIG. 5, where more than one is employed, they will usually be paralleland/or series). The slurry undergoes thickening in the hydrocyclone(s)228 resulting in a thickened underflow stream 230, e.g. up to about 30wt % solids, and an overflow stream 232 comprising solvent-based motherliquor in which terephthalic acid fines are suspended. The underflowstream is supplied to a pressure letdown vessel 234 in which thepressure of the thickened slurry is reduced to approximately 1 bara orbelow and is fed to the slurry receiving chamber of a rotary vacuumfilter 236 by means of which the terephthalic acid crystals are largelyseparated from mother liquor to produce a filter cake on the cylindricalfilter cloth of the vacuum filter. The filter cake is removed from thefilter cloth and supplied to a drier 237 to produce dried terephthalicacid crystals.

Solvent is recovered from the crystallisation process at various stages.The solvent and water-containing vapour flashed from vessels 222A, B isrecovered and passed through condensers 238A, B producing solvent ascondensate which is recycled via lines 240, 242 and 244 to the mixersection (see FIGS. 1, 2 and 4). Solvent is recovered as the overflowfrom the hydrocyclone and is recycled via line 232 and 244. Furthersolvent is recovered from solvent recovery section 246 and is recycledvia lines 248, 242 and 244. Solvent recovery section 246 may beconstituted by a distillation column (not shown), e.g. an azeotropicdistillation column, to which solvent and water-bearing feedstreamsarising from various sources within the production process are fed inorder to separate solvent from water, the water being passed to anaqueous effluent treatment system (not shown) via line 249. One suchfeedstream comprises the flash vapour stream 250 obtained from thecrystalliser vessels 222A, B following heat recovery in condensers 238A,B and 252. The vapour stream 250 will also contain residual oxygen andinerts such as nitrogen, carbon monoxide, carbon dioxide, solvent, waterand methyl acetate (the latter being present when air is used as thesource of oxygen). A second solvent/water feedstream 255 comprisingsolvent/water evaporated from a purge taken from the mother liquorrecycle stream is supplied to the distillation column . The gaseouscomponents are phase separated from the overheads produced in the courseof distillation and are supplied via line 254 to vent gas treatment(e.g. catalytic combustion, power recovery via an expander andscrubbing). The solvent recovered from the distillation column willusually be sufficiently clean to permit its use in washing of therecovered terephthalic acid. Thus, for instance, part of the solventstream 248 may be diverted via line 260 and used in washing of thefilter cake formed on the filter cloth of the rotary vacuum filter 236.If desired, the solvent may be used for countercurrent washing of thefilter cake on the filter cloth. Although a rotary vacuum filter isillustrated in FIG. 5, it will be appreciated that the solids-liquidseparation step may be carried out using other devices, e.g. a beltfilter. Instead of using solvent for washing, water may be used as thewash liquor.

FIG. 6 illustrates one form of solvent recovery from the crystallisationprocess in which energy is recovered by means of a condensing steamturbine. The product stream 300 from the reactor system is subjected tocrystallisation, two stages being shown. The crystallisation in eachvessel 302A, B takes place in the manner described with reference toFIG. 5 and results in a flash stream comprising steam and solventvapours. The flash from vessel 302A is passed to a condenser 304 inwhich heat is transferred to boiler feed water supplied via line 306thereby producing low pressure steam in line 308 for use in theproduction process. The condenser 304 serves to “knock back” solventwhich, because it is free of terephthalic acid, is passed via line 309to solvent recycle line 310 leading to the mixer/preheater associatedwithin the reactor rather than being transferred into the secondcrystalliser vessel 302B. The uncondensed, solvent-depleted flash ispassed via line 305 to the second crystalliser vessel 302B where itcombines with the flash from that vessel and is passed directly to afractional distillation/rectifier column 312. Alternatively the flashderived via line 305 can be fed directly to vent treatment as describedbelow. The flash vapours from the vessel 302B will be at pressure,typically of the order of 3 bara. The distillation column 312 willusually be arranged to receive the flash vapours from the lastcrystallisation stage where a series of crystallisation stages are used;however, we do not exclude the possibility of connecting thedistillation column with one of the crystallisation stages upstream ofthe final stage. The slurry from the vessel 302B is supplied to thesolids recovery process, depicted by reference 314, producing driedterephthalic acid crystals on line 316 and mother liquor recycle on line318.

The column 312 produces a bottoms product (line 320) comprising solventcontaining a small amount of water and a water-rich overheads product(line 322) containing some solvent. The solvent-rich bottoms product isrecycled to the reactor system via line 310 while the overheads productis processed to recover power by means of steam condensing turbine 324which may receive steam on input line 326, e.g derived from line 308,and lower pressure steam derived from the overheads product (line 322).The stream of residual organics, residual oxygen, nitrogen and steamforming the overheads product on line 322 is preheated by passagethrough heat exchanger 328 and by fired heater 330. It is next subjectedto catalytic combustion in unit 332 to destroy pollutants (mainlysolvent). The resulting high pressure, high temperature stream on line334 (typically at a temperature of the order of 450° C. followingcatalytic combustion) is cooled in heat exchanger 328 by heat transferto the incoming overheads stream on line 322 so as to adjust itstemperature for compatability with efficient operation of the condensingsteam turbine 324 to which it is supplied via line 336, i.e. to giveapproximately 12% wetness at the turbine outlet 338. The outlet streamfrom the turbine 338 is cooled in heat exchanger 342 and may, in part,be used as reflux (line 340) in the distillation column 312. Theremaining water recovered from the turbine 324 may be used elsewhere inthe production process or passed to effluent treatment, e.g. as boilerfeed water fed to condenser 304. In a modification of the distillationscheme of FIG. 6, all of the flash from the first crystalliser 302A maybe directly letdown into the second crystalliser 302B.

As mentioned previously, the slurry obtained in the course ofcrystallisation will be thin which means that concentration of theproduct relative to the mother liquor content is desirable beforeeffecting solids-liquid separation. FIG. 7 illustrates apparatus forsecuring concentration in the course of the crystallisation processusing an integrated crystallisation and concentration apparatus.Referring to FIG. 7, the crystallisation/concentrator section comprisesa Draft Tube Baffle (DTB) crystalliser. In FIG. 7, several sub-sectionsof plant (e.g. solvent dehydration, solids drying, purge treatment) andpumps and control valves are not shown to aid clarity of description.The omitted sub-sections of plant may be based on conventionalterephthalic acid production technology.

Oxygen supplied under pressure via line 400 and hot mother liquorrecycle supplied via line 402 are mixed in mixer 401 to pre-dissolve theoxygen and produce oxygenated mother liquor feed 404. At entry to plugflow reactor 410, feed 406 comprising fresh paraxylene, catalyst andacetic solvent is mixed with the oxygenated mother liquor feed 404 andreaction proceeds. The reaction exotherm results in increase intemperature through the reactor 410 and this, together with appropriateselection of the solvent:paraxylene ratio of the reaction medium,ensures that the terephthalic acid produced is maintained in solutionduring the reaction. The reactor product is fed forward via line 408 toa pressure (hence temperature) controlled DTB crystalliser 412. Onentering the central draft tube 413, the feed partially flashes. Becausethe two-phase flashing feed is less dense than the bulk mixture, acirculating flow is established within the crystalliser 412 (ifnecessary, the circulation is further enhanced via an up-pumpingagitator 414). Vapour is disengaged from the liquor surface and passesvia line 415 to a vapour condenser 416 that generates steam for processheating duties and power recovery. Most of the condensed solvent iscollected for recycle with the mother liquor via lines 418, 420 and 402but some is directed to solvent dehydration via line 422. The small gasvent stream from the condenser 416 (unreacted oxygen plus carbon oxidesand low levels of volatile organics) is passed forward to a vent gastreatment process via line 424.

The DTB crystalliser 412 has a settling zone 426 from which essentiallysolids free mother liquor is withdrawn (it will contain some finecrystals) for recycle on line 430. Product slurry is withdrawn from anelutriation leg 432 at higher slurry strength, e.g. up to about 30%solids by weight. In the elutriation leg 432, the crystals may be washedwith a countercurrent flow of clean solvent supplied via line 434. Thesolids-rich slurry is passed forward via line 436 for slurry cooling anddepressurising, for example by means of a single flash crystalliservessel 438 (or a number of crystallisers in series), with heat rejectionvia condenser 440 to cooling water and/or to heat recovery, theresulting non-condensibles being routed via line 442 and 424 to vent gastreatment and the condensate being routed via line 444 for recycle vialines 420 and 402. The cooled and depressurised slurry is then fedforward via line 435 to product separation section 446, for example arotary or belt filter or a centrifuge, operating at super, atmosphericor sub-atmospheric pressure. Wet solid product is recovered via line 448for product drying. Secondary mother liquor is recovered from theseparation section 446 via lines 450. A small purge of mother liquor istaken via line 452 to remove soluble impurities from the process. Thepurge is shown as taken from the product separation liquors but could betaken, for example, from the primary mother liquor stream 430.

Solvent (and optionally catalyst) recovered from the mother liquor purgestream 452, product drying, solvent dehydration and all crystallisercondenser condensate streams are combined, pre-heated (where necessaryby heater 454) and mixed with primary mother liquor recycle 430. Onmixing, any fines in the primary mother liquor will tend to dissolve (afines dissolution vessel may be provided to provide residence time forthis dissolution). Fines remaining undissolved will tend to go intosolution in the reactor as the solvent temperature increases.

As mentioned previously, the temperature profile through the reactorsystem, e.g. a single plug flow reactor, may be tailored torequirements, particularly to take into account constraints imposed bychemistry and chemical engineering considerations. For example, it willofter be desirable for the reactor inlet temperature to be as low aspossible provided that it is above the reaction initiation temperature.Advantageously, it will be at, or close to, the temperature at which theterephthalic acid is separated from the bulk of the mother liquor sincecooling or heating of recycled mother liquor is expensive due to thelarge flows involved. Also it has been found that product quality isstrongly influenced by the temperature at which solids-liquid separationis effected. These considerations imply the desirability of a reactorinlet temperature in the range 120-180° C., e.g. 140-170° C. The outlettemperature for an adiabatic (no external heating or cooling) reactor isrelated to the reactor inlet temperature and the solvent ratio. However,the outlet temperature is constrained by:

the need to minimise solvent and precursor burn (i.e. acetic acid andparaxylene burn to CO/CO₂) indicating the desirability of operating witha reactor outlet temperature below 230° C., e.g. 210° C.; and

the need to ensure that substantially all of the terephthalic acidproduced remains in solution by securing a suitable combination ofoutlet temperature and outlet solvent flow.

With an inlet temperature in the range 160-170° C. and an outlettemperature of no higher than 210° C., a simple adiabatic reactor wouldrequire a solvent:precursor ratio of greater than 100:1. This highsolvent ratio incurs significant capital and operating cost penaltiesaround the reactor, crystalliser(s), product recovery equipment andrecycle systems. Such high solvent:precursor ratios can be avoided byoperating the reactor system under conditions between adiabatic andisothermal by effecting removal of some of the heat of reaction, theheat so removed being used for example to raise steam for power recoveryand/or process heating duties.

One method of removing the heat of reaction while securing a desiredreactor outlet temperature is illustrated schematically in FIG. 8. Inthis embodiment, non-adiabatic/non-isothermal operation of the reactorsystem 510, supplied with reactant/solvent/recycle feeds 500, 502 issecured by internal cooling using one or more heat exchange means 512A,512B, 512C . . . through which a suitable coolant, e.g. boiler feedwater or mineral oil supplied by line 514, is circulated internallywithin the reactor system. As illustrated the heat exchangers are in theform of banks of tubes, the coolant flow being circulated through thetubes in co-current or counter-current relation with the flow ofreaction medium through the reactor system. Where the coolant compriseswater, the coolant may be removed as steam via line 518. The coolantused may alternatively be one of the streams employed in the process,e.g. the paraxylene feed, mother liquor recycle (before or after oxygenaddition and dissolution), so that the heat recovered is employed forinstance in raising the temperature of one or more of the feeds suppliedto the reactor inlet. Precipitation of terephthalic acid onto the heatexchange surfaces may be avoided by suitable choice of number, size andlocation of cooling tubes or coils, solvent:precursor ratio, solventoperating temperatures, steam raising temperature and flow pattern. Inthe latter context, the coolant may flow countercurrent and/orco-current relative to the reaction medium; however, co-current flow ispreferred. In FIG. 8, the reactor system may be constituted for exampleby a single plug flow reactor or it may comprise two or more plug flowreactors, one or more of which is provided with a heat exchanger asdescribed above to regulate temperature.

Although the invention as described with reference to the drawingsrefers to using paraxylene as the terephthalic acid precursor, it willbe appreciated that other precursors may be employed instead or inaddition to paraxylene, e.g. 4-tolualdehyde and 4-toluic acid.

EXAMPLES

Experimental work was carrried out using the plug flow reactor schemeillustrated in FIG. 9 . Vessel D302 is charged with a known quantity ofparaxylene in acetic acid/water solution. Vessel D301 is charged with aknown quantity of liquid catalyst in acetic acid/water solution. Airfrom supply AS is introduced into both D301 and D302, through dip pipes,by opening valve V20. The system pressure is set to ensure the desiredamount of oxygen (in excess of the stoichiometric paraxylenerequirement) goes into solution. Following oxygen dissolution, valve V21is opened and the differential pressure (delta P) controller DCP is setto establish a constant pressure between D301/D302 and the downstreamvessels. The fixed differential pressure fixes the reactor residencetime when liquid flow is established later.

Valves V23 and V25 are opened to cross-connect the two feed vessels D301and D302. Valves V27 and V28 are kept closed, preventing flow throughthe plug flow reactor in the form of Reaction Coil RC which is initiallyfilled with acetic acid. The vessels D301 and D302 and Reaction Coil areall immersed in an oil bath B which preheats the contents of vesselsD301 and D302 to the required reaction temperature. When D301/D302 areat temperature, reaction is started by opening valve V28 to establishflow through the Reaction Coil into an Off-Spec Vessel OSV andconsequent displacement of acetic acid from the Reaction Coil into theOff-Spec Vessel OSV. After a predetermined time, the product stream fromthe Reaction Coil is switched to the Sample Vessel SV by opening valveV27 and closing valve V28. Subsequently, the product stream from theReaction Coil RC is switched back to the Off-Spec Vessel OSV. At the endof the experiment, all vessels are cooled, vented via line AV, washedout and drained via drain lines D. The solid and liquid contents of theSample Vessel are recovered, weighed and analysed and the composition ofthe reaction solution leaving the Reaction Coil is back-calculated.

In Table 1, the concentrations of the reaction intermediates,paratolualdehyde (ptolald), paratoluic acid (ptol) and4-carboxybenzaldehyde (4-CBA) are reported for experiments where thereaction residence time was varied. At the small equipment scale used,the reactions run under quasi-isothermal conditions, close to the oilbath temperature of 210° C. throughout. The Examples clearly demonstratethe effect of residence time on intermediates concentrations. At 4.86minutes residence time, the paraxylene to reaction intermediates singlepass conversion is less than 0.5%. At 1.28 minutes residence time, theparaxylene to reaction intermediates single pass conversion is about16%. Significantly, however, paraxylene conversion to 4-CBA (theintermediate that tends to co-precipitate with the product, terephthalicacid) is of the order of 1% or below throughout.

TABLE 1 Plug Flow Reactor-Oxidation Results In all experiments thefollowing parameters were fixed (all compositions are w/w); Solventwater 5%, acetic acid 95% Paraxylene 0.5% w/w (200:1 solvent ratio)Catalyst Co 632 ppm, Mn 632 ppm, Br 1264 ppm + Zr 96 ppm Oil BathTemperature 210° C. Residence ptolald in ptol in 4CBA in Time solutionsolution solution Ex (min) (ppm w/w) (ppm w/w) (ppm w/w) 1 1.28 228 68776 2 1.78 55 411 51 3 2.28 132 312 42 4 2.31 99 192 38 5 3.29 15 82 6 64.86 1.7 27 <0.1

2. Crystallisation/Hot Filtration Experiments

A solution of 2% w/w terephthalic acid (TA), 125 ppm 4-CBA, 175 ppm ptoland other oxidation intermediates in 5% w/w water, 95% w/w acetic acidsolvent is prepared at elevated temperature (210° C.) and at a pressureto maintain a liquid phase. The solution is passed, continuously,through a pressure reducing valve into a crystalliser vessel whosepressure and temperature is controlled such that the TA is precipitatedfrom solution. The slurry produced in the crystalliser is passed forwardto further crystallisation vessels in which the pressure and temperatureare reduced to ambient conditions and further TA is precipitated.

During the course of the experiment, crystals from the firstcrystalliser (Hot Filtered TA) are recovered and are analysed for 4-CBAand paratoluic acid (ptol) content and median particle size (using aCoulter LS230 Laser Diffraction psd analyser). Crystals from thedownstream vessels (Cold Filtered TA) are also recovered and analysedfor reference purposes.

In Table 2, the Hot Filtered TA/4-CBA contents and median particle sizesare reported for experiments where the first crystalliser temperature,residence time and stirrer speed were varied. For reference, oneanalysis of Cold Filtered TA is also included. Examples 7, 8 and 9 showthat, in the Hot Filtered TA, 4-CBA and ptol contents fall as thefiltration temperature is reduced from 196 to 148° C. The data alsoshows that median particle size increases with reducing temperature. Ina separate experiment, Examples 10 and 11 show that, in the Hot FilteredTA, reduction in filtration temperature from 151 to 126° C. causes 4CBAlevel to increase, while ptol level and median particle size reduce.

When viewed together, Examples 7 through 11 indicate an optimum firstcrystalliser temperature, with respect to both intermediatesincorporation and median particle size, in the region 150+/−25° C.,especially 140 to 170° C.

Examples 12 and 13 show that increasing first crystalliser residencetime from 9 to 18 minutes benefits both intermediates incorporation andmedian particle size. Examples 14 and 15, when viewed alongside Example10, show that increasing first crystalliser agitator speed, from 270 to1000 rpm, does not have a strong influence on median particle size butreduces intermediates incorporation.

TABLE 2 Crystallisation/Hot Filtration Results in all experiments thefollowing parameters were fixed (all compositions are w/w); Solventwater 5%, acetic acid 95% Feed Solution Aromatics TA 2%, 4CBA 125 ppm,ptol 175 ppm Feed Solution Temperature 210° C. First Cryst. First Cryst.First Cryst. 4CBA ptol Median Res. Time Stirrer Speed Temp ContentContent Particle Size Ex (min) (rpm) (° C.) (ppm) (ppm) (micron)  7 121,000 196 2,360 345  59  8 12 1,000 176 1,040 218 114  9 12 1,000 148  670  89 134 10 18 1,500 151   710 138  96 11 18 1,500 126 1,060 117 86 12 18 1,000 173   980 150 106 13  9 1,000 179 1,140 217  96 14 12  270 152   930 123 139 15 12   500 150   790 106 135 Ref. 12 1,000 1482,340 281 102 (Cold Filter) (Cold Filter) (Cold Filter)

What is claimed is:
 1. A process for the production of terephthalic acidby the catalytic liquid phase oxidation of a precursor of terephthalicacid with oxygen in a reaction medium containing the precursor and analiphatic monocarboxylic acid solvent at a solvent:precursor ratio of atleast 30:1 which process comprises introducing the reactants into areaction zone which comprises at least one continuously stirred tankreactor in series with at least one plug flow reactor under conditionsof temperature and pressure whereby a continuous plug flow reactionregime is maintained and substantially all of the terephthalic acidproduced in the oxidation reaction remains in solution during thereaction.
 2. A process as claimed in claim 1 in which the oxidationreaction is carried out with substantially all of the oxygen dissolvedin the reaction medium.
 3. A process as claimed in claim 2 in which thereaction medium is produced by combining at least two separate liquidphase components and at least part of the oxygen is added to anddissolved in one or more of said liquid phase components before suchcomponents are combined to form the reaction medium.
 4. A process asclaimed in claim 2 or 3 in which oxygen is added to and dissolved in amother liquor recycle stream recovered from the reaction mediumfollowing completion of the reaction.
 5. A process as claim in claim 4in which the oxygen is introduced into the reaction in the form ofmolecular oxygen, as air or in an oxygen-containing gas.